Process for preparing a bio-diesel

ABSTRACT

The present invention relates to a process for preparing a bio-diesel, comprising the steps of, in the presence of an additional free fatty acid source, reacting a raw oil-fat with C 1 -C 6  monohydric alcohol in a reactor, and separating fatty acid esters from the reacted materials, so as to produce the bio-diesel, wherein the amount of the free fatty acid in the free fatty acid source ranges from 2-100 wt % and is higher than the amount of the free fatty acid in the raw fat-oil. The present process can increase the fatty acid ester yield and purity of raw oil-fats having a low reaction activity, and has a high adaptability to raw materials.

TECHNICAL FIELD

The present invention relates to a process for preparing a bio-diesel byreacting an oil-fat with a monohydric alcohol.

BACKGROUND OF THE INVENTION

Bio-diesel may be prepared by transesterification of an oil-fat with amonohydric alcohol. Besides fatty acid esters, the products of thetransesterification may include monoglycerides, diglycerides, glycerolby-products, as well as the unreacted alcohols and crude oil-fat.Bio-diesel primarily comprises fatty acid esters, and possibly othertrace substances such as monoglycerides, diglycerides, glycerol and thelike. In the prior art, there are the acid catalysis method, basecatalysis method, enzyme catalysis method and supercritical method forthe preparation of bio-diesel.

CN1473907A discloses using as raw materials the heels from the refiningof vegetable oils and the edible recovered oil, carrying out theproduction procedures comprising removing impurities by acidification,continuously dehydrating, esterifying, stratifying, and distilling underreduced pressure, and the catalyst used in the process is formed bycomplex formulation of inorganic and organic acids such as sulfuricacid, hydrochloric acid, p-toluene sulfonic acid, dodecylbenzenesulfonic acid, naphthalene sulfonic acid and the like. The continuousvacuum dehydration is carried out to a water content of less than 0.2%at a pressure of 0.08-0.09 MPa and a temperature of 60-95° C.Additionally, the catalyst is added in an amount of 1-3% in theesterification step; the esterification temperature ranges from 60-80°C.; and the reaction lasts 6 hours. After reaction, the product isneutralized to remove the catalyst, then stratified to remove water, anddistilled under reduced pressure to obtain a bio-diesel. The problems ofsaid acid catalysis include slow reaction rate, massive spent acids, andenvironmental pollution.

DE3444893 discloses a process, wherein an inorganic acid is used as thecatalyst; free fatty acids and alcohols are esterified at normalpressure and a temperature of from 50-120° C.; oils are pre-esterifiedand transesterified in the presence of an alkali metal catalyst.However, the residual inorganic acid catalyst will be neutralized withthe alkali, so as to increase the amount of the alkali metal catalyst.Moreover, the pre-esterification will lengthen the processing process,increase the equipment investment, greatly enhance the energyconsumption and incur a great loss of the materials. Moreover, the basiccatalyst needs to be removed from the product, and a great deal of wastewater will be produced.

CN1472280A discloses a process for preparing a bio-diesel, wherein fattyacid esters are used as the acyl receptor, and organisms are catalyzedfor interesterification in the presence of a bio-enzyme. However, thepresence of an enzyme catalyst has the disadvantages of long reactiontime, low efficiency, high enzyme price, and a high possibility ofinactivation in high purity methanol.

CN1142993C discloses a process for preparing fatty acid esters by usingan oil-fat and alcohol in the absence of a catalyst and under thecondition that either of said oil-fat and alcohol is in a supercriticalstate. The process is carried out using a batch kettle reaction, and isunfavorable to a large industry-scale production

CN1111591C discloses a process for preparing fatty acid esters bycontinuously reacting an oil-fat with a monohydric alcohol at atemperature of 270 to 280° C. and a pressure of 11-12 Mpa in a tubularreactor. However, the yield of fatty acid methyl ester is only 55-60%.

From the prior art above, it can be found that there are the problems oflower yield of bio-diesel and lower raw material processing capacity inthe preparation of a bio-diesel by medium and high pressure methods.

SUMMARY OF THE INVENTION

The present invention relates to a process for preparing a bio-diesel,comprising the steps of, in the presence of an additional free fattyacid source, reacting a raw oil-fat with C₁-C₆ monohydric alcohol in areactor, and separating fatty acid esters from the reacted materials, soas to produce the bio-diesel, wherein the amount of the free fatty acidin the free fatty acid source ranges from 2-100 wt % and is higher thanthe amount of the free fatty acid in the raw fat-oil. Said process canimprove the bio-diesel yield.

DETAILED DESCRIPTION OF THE INVENTION

The present invention relates to a process for preparing a bio-diesel,comprising the steps of, in the presence of an additional free fattyacid source, reacting a raw oil-fat with C₁-C₆ monohydric alcohol in areactor, and separating fatty acid esters from the reacted materials, soas to produce the bio-diesel, wherein the amount of the free fatty acidin the free fatty acid source ranges from 2-100 wt % and is higher thanthe amount of the free fatty acid in the raw fat-oil.

Said term oil-fat mentioned in, for example, the expression “a rawoil-fat” and “an oil-fat having a high acid number (or high acid numberoil-fat)” is generally known in the art and is a general designation ofoils and fats. The primary components thereof are fatty acidtriglycerides. Generally, an oil-fat in a liquid state at normaltemperature is termed as oil, and an oil-fat in a solid or semi-solidstate at normal temperature is termed as fat. Said oil-fat comprisesvegetable oils and animal oils, and further oils from microorganisms,algae and the like, and even crude oils, waste oil-fat and degenerativeoil-fat, and the like, wherein said crude oils are the oil-fat which isnot refined or fails to satisfy the product standard after refinement.The refining process includes, but is not limited to, degumming, alkalirefining, dephosphoration, decolorization, deodorization and the like.The oil-fat may also comprise unsaponifiable matters in a relativelyhigh content. Examples of vegetable oils comprise, but are not limitedto, soybean oil, rapeseed oil, peanut oil, sunflower seed oil, palm oil,cocoanut oil, and aliphatic group-containing substances from fruits,stems, leaves, limbs and roots derived from various agricultural cropsand wild plants (including a tall oil produced during the paper making).Examples of animal oil-fat include, but are not limited to, lard oil,beef tallow, mutton tallow, fish oil and the like. Said oil-fat may bethe mixture of two or more oil-fats.

The amount of the free fatty acid in said raw oil-fat is less than 50 wt%, preferably less than 30 wt %, more preferably less than 20 wt %. Inone embodiment, the raw oil-fat comprises palm oil. In anotherembodiment, the raw oil-fat is a waste oil-fat.

Said C₁-C₆ monohydric alcohol is a monohydric fatty alcohol having from1 to 6 carbon atoms, which may be a saturated or unsaturated alcohol.Examples of the monohydric alcohol include, but are not limited to,methanol, ethanol, n-propanol, isopropanol, allyl alcohol, butanol suchas n-butanol, isobutanol, amyl alcohol such as n-amyl alcohol and thelike. A single alcohol, or a mixture of two or more alcohols may beused. Said monohydric alcohol is preferably selected from methanol andethanol, especially methanol. The molar ratio of C₁-C₆ monohydricalcohol to the raw oil-fat may range from 3 to 60:1, preferably from 4to 12:1.

Said free fatty acid source can be a free fatty acid which may besaturated or unsaturated, preferably C₁₀-C₂₄ saturated or unsaturatedfatty acid, more preferably C₁₂-C₁₈ unsaturated fatty acid which may hasone or more double-bonds, preferably one double-bond. Examples of thefree fatty acid comprise, but not limited to, tetracosanoic acid,docosanoic acid, eicosanic acid, nonadecanoic acid, stearic acid,heptadecoic acid, palmitic acid, pentadecanoic acid, myristic acid,tridecanoic acid, lauric acid, undecanoic acid, decanoic acid,docosenoic acid, arachidonic acid, oleic acid, linolenic acid, linoleicacid, undecenoic acid, etc. A specially preferred free fatty acid isoleic acid. Said free fatty acid is present in an amount of 1-50 wt %,preferably 2-40 wt %, relative to the weight of the raw oil-fat.

Said free fatty acid source can also be the oil-fat having a high acidnumber, for example crude oil, waste oil-fat, or the like. The amount ofthe free fatty acid in the oil-fat having a high acid number is morethan 2 wt %, preferably 5 to <100 wt %, more preferably 10 to 60 wt %,and higher than the amount of the free fatty acid in the raw oil-fat.The weight ratio between the raw oil-fat and the oil-fat having a highacid number ranges from 1:0.02 to 50, preferably 1:0.04 to 20, morepreferably 1:0.06 to 10. The category of the oil-fat having a high acidnumber can be same or different from that of the raw oil-fat.

The reaction between the raw oil-fat and C₁-C₆ monohydric alcohol in thepresence of an additional free fatty acid source can be optionallycarried out under the condition that an alkaline compound is used as acatalyst. Said alkaline compound can be selected from, e.g. hydroxides,alcoholates, oxides, carbonates, bicarbonates and aliphatic carboxylateof the Groups IA and IIA elements in the periodic table, preferablyhydroxides, alcoholates, oxides, carbonates, bicarbonates and C₁₂-C₂₄fatty acid salts of sodium, potassium, magnesium, calcium and barium,more preferably hydroxides, oxides, alcoholates and C₁₂-C₂₄ fatty acidsalts of sodium and potassium. Said alkaline compound is added in anamount of 0.005-0.3 wt %, preferably 0.008-0.2 wt %, relative to theweight of the oil-fat.

In the process of the present invention, a tubular reactor may be used.The reactor may be provided with the oil-fat and alcohols separately orafter being pre-mixed. The materials may be preheated by a pre-heaterbefore being fed into the reactor, or directly fed into the reactor. Ifthe materials are directly fed into the reactor, the reaction functionsas both a pre-heater and a reactor. If a pre-heater is used, the oil-fatand alcohol may be pre-heated respectively or pre-heated together afterthey are mixed. The reaction temperature ranges from 200 to 320° C.,especially from 230 to 280° C.; the reaction pressure ranges from 5 to12 MPa, especially from 6 to 10 MPa, and a relatively low pressure (e.g.5 to 7.5 MPa) can also achieve the objective of the present invention;the liquid hourly space velocity of the oils ranges from 0.1 to 10 h⁻¹,especially from 0.5 to 6 h⁻¹, more especially from 1 to 3 h⁻¹. Said oilscomprise the raw oil-fat and additional free fatty acid source.

In the process of the present invention, the separation of fatty acidesters comprises the steps of

-   -   (A) separating the mixed ester phase and the glycerol phase        formed in the reacted materials, and subsequently evaporating        monohydric alcohols respectively from said mixed ester phase and        optionally from the glycerol phase, or evaporating monohydric        alcohols from the reacted materials before separating the mixed        ester phase and the glycerol phase formed in the reacted        materials; and    -   (B) distilling or rectifying the mixed ester phase processed in        step (A), or water-washing the mixed ester phase processed in        step (A) and separating the ester phase formed after washing        from the aqueous phase and collecting said ester phase, to        obtain high purity fatty acid esters, and optionally distilling        the glycerol phase processed in step (A) to obtain glycerol.

In step (A) above, monohydric alcohols may be evaporated byrectification or flash distillation under the condition that thetemperature at the column bottom is less than 150° C., and the pressuremay be normal pressure, vacuum, or greater than one atmosphericpressure.

In step (A) above, the mixed ester phase and glycerol phase may beseparated by deposition or via a fiber bundle separator. Rapidseparation via a fiber bundle separator is preferred. Said fiber bundleseparator consists of a separating cylinder and a receiving tank,wherein the separating cylinder is furnished with fiber bundlesconsisting of stainless steel wires. The mixture of the mixed esterphase and the glycerol phase firstly passes through the separatingcylinder and then is fed into the receiving tank for stratification, soas to achieve the separation of the mixture. The separation is carriedout at a temperature of from 20 to 200° C., preferably from 40 to 100°C., at a pressure of greater than one atmospheric pressure or at normalpressure, e.g. from 0.1 to 0.5 MPa, preferably from 0.1 to 0.3 MPa, andat a space velocity of from 0.1 to 25 h⁻¹, especially from 1 to 10 h⁻¹,more preferably from 1 to 5 h⁻¹. In order to achieve better phaseseparation effect, the reacted materials which are heavily emulsifiedgenerally need to stand overnight if the deposition method is used. Afiber bundle separator can achieve a good separation effect in a veryshort time, so as to greatly enhance the separation rate and productionefficiency.

In one embodiment of step (B) above, the mixed ester phase processed instep (A) above is distilled or rectified to obtain high purity fattyacid esters, wherein the distillation or rectification of said mixedester phase may be carried out in a rectification column under reducedpressure or via a film evaporator. The mixed ester phase obtained instep (A) is fed into the reduced pressure rectification column, whereinthe column bottom pressure is less than 0.1 MPa, preferably less than0.01 MPa, more preferably less than 0.001 MPa; reflux may not occur, orthe reflux ratio is controlled to range from 0.01 to 10:1, preferablyfrom 0.1 to 2:1. The temperature of the column bottom or film evaporatorranges from 100 to 300° C., preferably from 170 to 280° C., morepreferably from 190 to 280° C. The optional distillation of the glycerolphase may be similarly carried out by rectification in a reducedpressure rectification column or via a film evaporator.

In another embodiment of step (B) above, the mixed ester phase processedin step (A) above is washed by water, and the ester phase formed afterwashing is separated from the aqueous phase and collected to obtain highpurity fatty acid esters. The water to be added during the washing isfrom 10 to 100 wt %, preferably from 20 to 80 wt % relative to theamount of the mixed ester phase; the temperature of water ranges from 25to 100° C., preferably from 40 to 80° C. Washing may be carried out onceor more times. If the ester phase obtained in step (B) has a relativelyhigh acid number, an alkaline substance may be added into washing water.One or more alkaline substances selected from the group consisting ofsodium carbonate, sodium bicarbonate, potassium carbonate, potassiumbicarbonate, sodium hydroxide and potassium hydroxide may be added inthe form of an aqueous solution for alkaline washing, wherein thealkaline substance in the aqueous solution is in a concentration of from5 to 40 wt %, preferably from 5 to 20 wt %. The washed mixture may bere-separated into the ester phase and the aqueous phase by e.g.deposition, preferably via a fiber bundle separator, at a temperature offrom 20 to 150° C., preferably from 40 to 100° C., at a pressure ofgreater than one atmospheric pressure or at normal pressure, and at aspace velocity of from 0.1 to 25 h⁻¹, especially from 1 to 10 h⁻¹, morepreferably from 1 to 5 h⁻¹.

The process of the present invention may further comprise step (C):separating monoglycerides and diglycerides from the mixed ester phaseresidues (i.e. residual liquid in the column bottom) distilled orrectified in step (B) by using a secondary molecular rectification, orevaporating monoglycerides and diglycerides from the mixed ester phaseresidues distilled or rectified in step (B) by using a primary molecularrectification. The evaporated monoglycerides and/or diglycerides may berecycled as required to the reactor inlet for a second reaction. Morespecifically, if the fraction having a higher monoglyceride content isdesired, a secondary molecular rectification may be used. The residualliquid obtained from step (B) in the column bottom is fed into themolecular rectification device. The monoglyceride fraction in a highercontent may be obtained at a pressure of less than 5 Pa, preferably lessthan 3 Pa, more preferably less than or equal to 1 Pa, and at a heatingsurface temperature of from 170 to 220° C., preferably 180 to 200° C.The fraction having a higher monoglyceride content may be used as an oillubricating additive, and the heavy fractions may be fed into thesecondary molecular rectification. At the pressure above and at aheating surface temperature of from 200 to 290° C., preferably 220 to250° C., diglycerides having a higher purity may be obtained. Thesemonoglycerides and diglycerides can be recycled as raw materials to thereactor inlet for second reaction. If the monoglyceride fraction in ahigher content is not necessary, monoglycerides and diglycerides may bedirectly evaporated by using a primary molecular rectification, and thenrecycled to the reactor inlet for second reaction. In addition, theheavy residue may be used as fuel. In order to achieve the object ofseparating more components, the continuous multistage (or multi-group)operation may be used for the molecular rectification.

According to the present process, the biodiesel yield of the raw oil-fathaving a low reaction activity is increased, the defect that the productobtained from the high acid number raw oil-fat has a high acid number isavoided, and the biodiesel yield resulting from the mixed oils as rawmaterials are significantly higher than the sum of the yields resultingfrom the respective low activity oils and high acid number oil-fat whenthey are reacted alone. High purity of the fatty acid ester can also beachieved together with a strong raw material-processing capability, andthus the inventive process makes merit in industry applications. Theprocess of the present invention has a strong adaptability to rawmaterials. Even if the oil-fat has a very high acid number or evencontains non-saponifiable matters in a higher content, it can bedirectly processed without the multifarious pretreatment so as to reducethe energy consumption and equipment investment. In addition, theprocess of the present invention may effectively separatemonoglycerides, diglycerides and organic matters in non-refined oilshaving a relatively high boiling point, and may enable the components innon-refined oils which can become fatty acid methyl esters to beutilized to a maximum extent.

The present process is easy to be carried out, has substantially noliquid, gas, and solid waste, and is environmentally friendly.

EXAMPLES

The following examples are used to further explain the presentinvention, but the present invention is not limited to these examples.The raw materials used below are commercially available or easilyproduced according to the common technology in the art.

The bio-diesel yield stated in the examples can be calculated from theratio of the bio-diesel weight to the weight sum of the raw oil-fat andfree fatty acid source; the purity of fatty acid methyl ester can becalculated from the ratio of the fatty acid methyl ester weight to thebio-diesel weight.

Comparison Example 1

A refined soybean oil having 0.7 wt % of free fatty acid wascontinuously fed as raw materials into a tubular reactor at an oilliquid hourly space velocity of 1.2 h⁻¹ and methanol:oil molar ratio of5. The reaction temperature was 280° C. and the pressure was 9.5 MPa.Methanol and glycerol were separated from the raw product of reaction,which then was vacuum-rectified to evaporate the bio-diesel. The yieldof the bio-diesel was 47 wt %.

Comparison Example 2

A rapeseed oil having 0.45 wt % of free acid was continuously fed as rawmaterials into a tubular reactor at an oil liquid hourly space velocityof 1.2 h⁻¹ and methanol:oil molar ratio of 7. The reaction temperaturewas 272° C. and the pressure was 8 MPa. Methanol and glycerol wereseparated from the raw product of reaction, which then wasvacuum-rectified to evaporate the bio-diesel. The yield of thebio-diesel was 61 wt %.

Example 1

A crude soybean oil having 1.2 wt % non-saponifiable matters and 36 wt %free fatty acid was mixed with the rapeseed oil in comparison example 2in a mixing ratio of 0.3:1. The mixture was continuously fed into atubular reactor at an oil liquid hourly space velocity of 1.2 h⁻¹ andmethanol:oil molar ratio of 7. The reaction temperature was 272° C. andthe pressure was 8 MPa. Methanol and glycerol were separated from theraw product of reaction, which then was vacuum-rectified to evaporatethe bio-diesel. The yield of the bio-diesel was 92.2 wt %, and the acidnumber of the bio-diesel was 3.5 mgKOH/g, wherein the fatty acid methylester had a purity of 98 wt %. The residual liquid in the column bottomcould be recycled as raw material to the reactor inlet for a secondreaction. The components in the raw materials which could become fattyacid methyl esters were almost converted to the desired product.

Under the same conditions, above crude soybean oil having 36 wt % freefatty acid was used as raw materials, and a bio-diesel was produced in ayield of 87.2 wt % and had an acid number of 9.0 mgKOH/g.

Example 2

A cotton seed oil having 1.5 wt % non-saponifiable matters and 26 wt %free fatty acid was mixed with the refined soybean oil in comparisonexample 1 in a mixing ratio of 1:1. The mixture was fed into a tubularreactor at an oil liquid hourly space velocity of 1.2 h⁻¹ andmethanol:oil molar ratio of 5. The reaction temperature was 272° C. andthe pressure was 8.5 MPa. Methanol and glycerol were separated from theraw product of reaction, which then was vacuum-rectified to evaporatethe bio-diesel. The yield of the bio-diesel was 91 wt %, and the acidnumber of the bio-diesel was 3.3 mgKOH/g, wherein the fatty acid methylester had a purity of 97.5 wt %, and the free glycerol content was 0.2wt %. To the ester phase, 5% sodium carbonate solution at 40° C. wasadded for washing. The washed mixture was fed into a fiber bundleseparator at a temperature of 40° C. and a liquid hourly space velocityof 10 h⁻¹. The ester phase and the aqueous phase were separated. Theacid number of the ester phase was 0.27 mgKOH/g and the content of thefree glycerol was 0.018 wt %. The residual liquid in the column bottomcould be recycled as raw material to the reactor inlet for a secondreaction.

Under the same conditions, above cotton seed oil having 26 wt % freefatty acid was used as raw materials, and a bio-diesel was produced in ayield of 86 wt % and had an acid number of 7.2 mgKOH/g.

Example 3

A waste fat-oil having 14.3 wt % free fatty acid was mixed with asoybean oil and rapeseed oil in a mixing ratio of 1:1:1. The mixture wascontinuously fed into a tubular reactor at an oil liquid hourly spacevelocity of 1.2 h⁻¹ and methanol:oil molar ratio of 5. The fatty acidmethyl ester was produced under a reaction temperature of 320° C. and apressure of 9 MPa. The crude product flowing out from the reactor wasthen fed into a flash column to remove methanol at a temperature of lessthan 150° C., and recycle and reuse the methanol. The residual materialswere fed into a separator comprising fiber bundles, and an ester phasewas separated out at a temperature of 52° C. and a liquid hourly spacevelocity of 5 h⁻¹, and the ester phase was fed into a vacuumrectification column. The bio-diesel was obtained at the column topunder the conditions comprising a vacuum degree of 8 mmHg, a columnbottom temperature of 280° C. and no reflux. The yield of the bio-dieselwas 83.9 wt %, the acid number was less than 2.0 mgKOH/g, and the fattyacid methyl ester purity was 98.5 wt %. The residual liquid in thecolumn bottom was fed into the molecular rectification device to obtaina light fraction having a relatively high content of monoglyceride at aresidual pressure of from 5 to 6 Pa and a heating surface temperature of190° C. The remaining materials having a high boiling point were fedinto a secondary molecular rectification to obtain a light fraction at aresidual pressure of 2 Pa and a heating surface temperature of 240-244°C., wherein said light fraction as the raw materials could be recycledto the reactor inlet for the second reaction. The components in the rawmaterials which could become fatty acid methyl esters were almostconverted to the desired product.

Under the same conditions, above waste oil-fat having 14.3 wt % freefatty acid was used as raw materials, and the bio-diesel was produced ina yield of 84.2 wt % and had an acid number of 4.7 mgKOH/g.

Under the same conditions, the bio-diesels were produced in yields of 69wt % and 76 wt % when a soybean oil and rapeseed oil were used as rawmaterials respectively.

Example 4

A waste fat-oil having 50 wt % free fatty acid was mixed with therapeseed oil in comparison example 2 in a mixing ratio of 0.2:1. Themixture was continuously fed into a tubular reactor at an oil liquidhourly space velocity of 1.2 h⁻¹ and methanol:oil molar ratio of 6. Thereaction temperature was 272° C. and the pressure was 7.2 MPa. The crudeproduct flowing out from the reactor was then depressurized to 0.1-0.13MPa, then the liquid was entered into a fiber bundle separator at atemperature of 30° C. and a liquid liquid hourly space velocity of 5 h⁻¹to separate out an ester phase and a glycerol phase. The ester phase andglycerol phase were fed into the respective flash columns to removemethanol at a temperature less than 150° C., and recycle and reuse themethanol. After methanol was evaporated, the glycerol was separated andthe mixed ester was fed into a film evaporator to evaporate a bio-dieselat a vacuum degree of 8 mmHg and a temperature of 252° C. The yield ofthe bio-diesel was 93 wt %, the acid number was 4.2 mgKOH/g, and thefatty acid methyl ester purity was 97.3 wt %. The residual liquid in thecolumn bottom was recycled as raw materials to a reactor inlet for asecond reaction. The components in the raw materials, which could becomefatty acid methyl esters, were almost converted to the desired product.

Under the same conditions, the bio-diesel was produced in a yield of 90wt % when above waste oil-fat was used as raw materials.

Example 5

A raw rapeseed oil having 29.5 wt % free fatty acid was mixed with therapeseed oil in comparison example 2 in a mixing ratio of 40:1. Themixture was continuously fed into a tubular reactor at an oil liquidhourly space velocity of 1.2 h⁻¹ and methanol:oil molar ratio of 6. Thereaction temperature was 272° C. and the pressure was 9.5 MPa. The crudeproduct flowing out from the reactor was depressurized, and then stoodfor deposition to separate out a mixed ester phase and glycerol phase.The mixed ester phase and glycerol phase were fed into the respectiveflash columns to continuously flash evaporate methanol respectively at atemperature less than 150° C. The mixed ester phase in which methanolhad been evaporated was fed into a vacuum rectification column toevaporate a bio-diesel at the column top at a vacuum degree of 8 mmHg, acolumn bottom temperature of 252-255° C. and a reflux ratio of 1:1. Theyield of the bio-diesel was 86 wt %, and the fatty acid methyl esterpurity was up to 99.5 wt %. The residual liquid in the column bottom wasfed into a molecular rectification device to evaporate a light fractionat a residual pressure of 1 Pa and a heating surface temperature of 250°C. Said light fraction might be recycled as raw materials to a reactorinlet for a second reaction. The components in the raw materials, whichcould become fatty acid methyl esters, were almost converted to theproduct.

Example 6

As raw materials, a cotton seed oil having 5 wt % non-saponifiablematters and 15 wt % free fatty acid, in which 10 wt % oleic acid wasadded, was continuously fed into a tubular reactor at an oil liquidhourly space velocity of 1.2 h⁻¹ and methanol:oil molar ratio of 4.5.The reaction temperature was 272° C. and the pressure was 8 MPa.Methanol and glycerol were separated from the raw product of reaction,which then was vacuum-rectified to evaporate the bio-diesel. The yieldof the bio-diesel was 91.3 wt %, and the acid number of the bio-dieselwas 3.5 mgKOH/g, and the fatty acid methyl ester had a purity of 97.7 wt%.

The residual liquid in the bottom of the rectification column containedfatty acid methyl ester, monoglyceride, and diglyceride, could berecycled as raw materials to a reactor inlet, and reused after beingmixed with fresh raw materials.

Example 7

As raw materials, a waste oil-fat having 4.5 wt % free fatty acid and1.2 wt % non-saponifiable matters, in which 20 wt % oleic acid wasadded, was continuously fed into a tubular reactor at an oil liquidhourly space velocity of 1.2 h⁻¹ and methanol:oil molar ratio of 5. Thereaction temperature was 272° C. and the pressure was 7.4 MPa. The crudeproduct flowing out from the reactor was depressurized, and then fedinto a rectification column to remove methanol at a temperature lessthan 150° C., and recycle and reuse the methanol. The remainingmaterials were stood for deposition to separate out an ester phase andglycerol phase. The ester phase was fed into a vacuum rectificationcolumn to evaporate a bio-diesel at the column top at a vacuum degree of8 mmHg, a column bottom temperature of 251-255° C. and a reflux ratio of1:1. The yield of the bio-diesel was 78 wt %. The residual liquid in thecolumn bottom was fed into a molecular rectification device to evaporatea light fraction at a residual pressure of 1 Pa and a heating surfacetemperature of 250° C. Said light fraction might be recycled as rawmaterials to a reactor inlet for a second reaction. The components inthe raw materials, which could become fatty acid methyl esters, werealmost converted to the product.

When the above same raw materials, reaction conditions and separationsteps were applied, except for no addition of oleic acid during thereaction, a bio-diesel yield of 45.9 wt % was resulted.

Example 8

As raw materials, a palm oil having 0.35 wt % free fatty acid, in which15 wt % oleic acid was added, was continuously fed into a tubularreactor at an oil liquid hourly space velocity of 1.2 h⁻¹ andmethanol:oil molar ratio of 8. The reaction temperature was 272° C. andthe pressure was 8 MPa. The crude product flowing out from the reactorwas depressurized to 0.1-0.13 MPa, and then fed into a fiber bedcontaining fiber bundles to separate out a mixed ester phase andglycerol phase at a temperature of 40° C. and a liquid liquid hourlyspace velocity of 7 h⁻¹. The mixed ester phase was fed into arectification column to remove the methanol at a temperature less than150° C. and recycle and reuse the methanol. The remaining materials werefed into a fiber bed containing fiber bundles to separate out glycerolat a temperature of 40° C. and a liquid hourly space velocity of 7 h⁻¹.The ester phase was vacuum-rectified at a vacuum degree of 5 mmHg and acolumn bottom temperature of 235-240° C. The yield of the bio-diesel was69.6 wt %.

The residual liquid in the bottom of the distillation column containedfatty acid methyl ester, monoglyceride, and diglyceride, could berecycled as raw materials to a reactor inlet, and reused after beingmixed with fresh raw materials. The components in the raw materials,which could become fatty acid methyl esters, were almost converted tothe product.

When the above same raw materials, reaction conditions and separationsteps were applied, except for no addition of oleic acid during thereaction, a bio-diesel yield of 34.9 wt % was resulted.

Example 9

As raw materials, a soybean oil having 2.5 wt % free fatty acid, inwhich 3 wt % free unsaturated C₁₆ fatty acid was added, was continuouslyfed into a tubular reactor at an oil liquid hourly space velocity of 1.2h⁻¹ and methanol:oil molar ratio of 15. The reaction temperature was300° C. and the pressure was 9 MPa. The crude product, from whichmethanol and glycerol had been separated, was vacuum-rectified toproduce the bio-diesel in a yield of 72 wt %.

The residual liquid in the bottom of the rectification column containedfatty acid methyl ester, monoglyceride, and diglyceride, could berecycled as raw materials to a reactor inlet, and reused after beingmixed with fresh raw materials.

When the above same raw materials, reaction conditions and separationsteps were applied, except for no addition of free unsaturated C₁₆ fattyacid during the reaction, a bio-diesel yield of 55 wt % was resulted.

It should be understood that various changes and modifications to thepresent invention can be made by those skilled in the art withoutdeparting the inventive scope so as to be applicable for the variousobjectives and conditions. Therefore, these changes and modificationsshould be suitably and reasonably encompassed in all the equivalentscopes of the accompanying claims.

1. A process for preparing a bio-diesel, comprising the steps of, in thepresence of an additional free fatty acid source, reacting a raw oil-fatwith C₁-C₆ monohydric alcohol in a reactor, and separating fatty acidesters from the reacted materials, so as to produce the bio-diesel,wherein the amount of the free fatty acid in the free fatty acid sourceranges from 2-100 wt % and is higher than the amount of the free fattyacid in the raw fat-oil.
 2. The process according to claim 1,characterized in that said free fatty acid source is a free fatty acid.3. The process according to claim 2, characterized in that said freefatty acid is a C₁₀-C₂₄ saturated or unsaturated fatty acid.
 4. Theprocess according to claim 2, characterized in that said free fatty acidis a C₁₂-C₁₈ unsaturated fatty acid.
 5. The process according to claim2, characterized in that said free fatty acid is oleic acid.
 6. Theprocess according to claim 2, characterized in that said free fatty acidis present in an amount of 1-50 wt %, relative to the weight of the rawoil-fat.
 7. The process according to claim 2, characterized in that saidfree fatty acid is present in an amount of 2-40 wt %, relative to theweight of the raw oil-fat.
 8. The process according to claim 1,characterized in that said free fatty acid source is an oil-fat having ahigh acid number wherein the amount of the free fatty acid is in anamount of more than 2 wt %.
 9. The process according to claim 8,characterized in that the amount of the free fatty acid in said oil-fathaving a high acid number is 5 to <100 wt %.
 10. The process accordingto claim 8, characterized in that the amount of the free fatty acid insaid oil-fat having a high acid number is 10 to 60 wt %.
 11. The processaccording to claim 8, characterized in that said oil-fat having a highacid number is a raw oil or waste oil-fat.
 12. The process according toclaim 8, characterized in that the weight ratio between the raw oil-fatand the oil-fat having a high acid number ranges from 1:0.02 to
 50. 13.The process according to claim 8, characterized in that the weight ratiobetween the raw oil-fat and the oil-fat having a high acid number rangesfrom 1:0.04 to
 20. 14. The process according to claim 8, characterizedin that the weight ratio between the raw oil-fat and the oil-fat havinga high acid number ranges from 1:0.06 to
 10. 15. The process accordingto claim 1, characterized in that the amount of the free fatty acid insaid raw oil-fat is less than 50 wt %.
 16. The process according toclaim 15, characterized in that the amount of the free fatty acid insaid raw oil-fat is less than 30 wt %.
 17. The process according toclaim 16, characterized in that the amount of the free fatty acid insaid raw oil-fat is less than 20 wt %.
 18. The process according toclaim 1, characterized in that said raw oil-fat comprises palm oil. 19.The process according to claim 1, characterized in that said raw oil-fatis a waste oil-fat.
 20. The process according to claim 1, characterizedin that said C₁-C₆ monohydric alcohol is methanol or ethanol.
 21. Theprocess according to claim 1, characterized in that the molar ratio ofC₁-C₆ monohydric alcohol to the raw oil-fat ranges from 3 to 60:1. 22.The process according to claim 1, characterized in that the molar ratioof C₁-C₆ monohydric alcohol to the raw oil-fat ranges from 4 to 12:1.23. The process according to claim 1, characterized in that the reactionbetween the raw oil-fat and C₁-C₆ monohydric alcohol in the presence ofan additional free fatty acid source is carried out under the conditionthat an alkaline compound is used as a catalyst.
 24. The processaccording to claim 1, characterized in that said reactor is a tubularreactor
 25. The process according to claim 1, characterized in that thereaction temperature ranges from 200 to 320° C.
 26. The processaccording to claim 1, characterized in that the reaction temperatureranges from 230 to 280° C.
 27. The process according to claim 1,characterized in that the reaction pressure ranges from 5 to 12 MPa. 28.The process according to claim 1, characterized in that the reactionpressure ranges from 5 to 10 MPa.
 29. The process according to claim 1,characterized in that the reaction pressure ranges from 5 to 7.5 MPa.30. The process according to claim 1, characterized in that the totalliquid hourly space velocity of the raw oil-fat and additional freefatty acid source ranges from 0.1 to 10 h⁻¹.
 31. The process accordingto claim 1, characterized in that the total liquid hourly space velocityof the raw oil-fat and additional free fatty acid source ranges from 0.5to 6 h⁻¹.
 32. The process according to claim 1, characterized in thatthe separation of fatty acid esters comprises the steps of (A)separating the mixed ester phase and the glycerol phase formed in thereacted materials, and subsequently evaporating monohydric alcoholsrespectively from said mixed ester phase and optionally from theglycerol phase, or evaporating monohydric alcohols from the reactedmaterials before separating the mixed ester phase and the glycerol phaseformed in the reacted materials; and (B) distilling or rectifying themixed ester phase processed in step (A), or water-washing the mixedester phase processed in step (A) and separating the ester phase formedafter washing from the aqueous phase and collecting said ester phase, toobtain fatty acid esters, and optionally distilling the glycerol phaseprocessed in step (A) to obtain glycerol.
 33. The process according toclaim 32, characterized in that the monohydric alcohol is evaporated instep (A) by flash distillation or rectification under the condition thatthe temperature at the column bottom is less than 150° C.
 34. Theprocess according to claim 32, characterized in that, in step (A), theseparation between the mixed ester phase and the glycerol phase iscarried out by deposition or via a fiber bundle separator.
 35. Theprocess according to claim 34, characterized in that said fiber bundleseparator consists of a separating cylinder and a receiving tank,wherein the separating cylinder is furnished with fiber bundlesconsisting of stainless steel wires, and the separation between themixed ester phase and the glycerol phase is carried out by passing themthrough the separating cylinder and then feeding into the receiving tankfor stratification.
 36. The process according to claim 32, characterizedin that, during the washing process in step (B), an alkaline matter isadded into the washing water.
 37. The process according to claim 32,characterized in further comprising step (C): evaporating monoglyceridesand diglycerides from the mixed ester phase residues distilled orrectified in step (B) by using a primary molecular rectification, orevaporating and separating monoglycerides and diglycerides from themixed ester phase residues distilled or rectified in step (B) by using asecondary molecular rectification, with or without a second reaction.